Catalytic reforming of straight-run or cracked naphtha fractions in the presence of added hydrogen in a multiple reactor fixed-bed system



4 Sheets-Sheet 1 R o T c I E R o N N02 REACTOR GAS TO FUEL FLA$H DRUM REFORMATE INTERHEATER DECKER FRACTIONS IN THE PRESENCE OF ADDED HYDROGEN IN A MULTIPLE REACTOR FIXED-BED SYSTEM CATALYTIC REFORMING 0F STRAIGHT-RUN OR CRACKED NAPHTHA RECYCLE GAS PREHEATER FUEL RECYCLE GAS COMPRESSOR July Filed Aug. 27, 1956 William H.Decker INVENTOR.

BYMLUI @WSMZZM ATTORNEYS July 5, 1960 w. H; DECKER 2,943,998

CATALYTIC REFORMING OF STRAIGHT-RUN 0R CRACKED NAPI-I'IHA FRACTIONS IN THE PRESENCE OF ADDED HYDROGEN v IN A MULTIPLE REACTOR FIXED-BED sys'rsm Filed Aug. 27, 1956 4 Sheets-Sheet 2 3g CATALYST sen VAPOR INLET [111]]1- DISTRIBUTION 20rd:

"CATALYST RETENTION ZONE FIG. 2

REHEAT INJECTION NOZZLES 31 42 mznuocoun:

conuzcnon VAPOR OUTLET DISTRIBUTION HEADERS FIG. 3

44 William H. Decke! INVENTOR.

BYM amvhzlfim ATTORNEYS y 5, 1930 w. H. DECKER 2,943,998

CATALYTIC REFORMING 0F STRAIGHT-RUN 0R CRACKED NAPHTHA FRACTIONS IN THE PRESENCE 0F ADDED HYDROGEN IN A MULTIPLE REACTOR FIXED-BED. SYSTEM Filed Aug. 27, 1956 v I 4 Sheets-Sheet 4 kLlDO'HA EOVdS O1. EHDLVUHdWBL NOLLDVEH d0 dIHSNOILV'IHB William H. Decker INVENTOR.

ATTORNEYS 2,943,998 E Patentedgrluly. 19.60

.United States atent i 2,943,998 csystem withra consequent reduction in capitalrequirement: fOl'iIhE reactor units and for catalyst 'myentory. CATALYTIC REFORMINGQF STRAIGHT-RUN. The, reforming; process :of my inventionihusinvolYes CRACKED NAPHTHA FRACTIONS IN THE PRES :basically the maintenance of-semi-isothermal temperature "ENCE 0F ADDED HYDROGEN IN A MULTIPLE 6 conditions in the first reactor of a'multi'ple reactonfixed- I REACTOR FIXED'BED SYSTEM "bed systern by limiting the amount of hydrogen-contain- Willi am H. Decker, Hazel Crest, 111., assignor toSinclair :i g i d i t e with. hydrocarbon vapors initially Re'fining (Iompany; New York, N.Y.,-a corporation of Lcharged toilthe'i first vreactoncatalyst bedand. separately Maine 1introducingadditionalpro-heated hydrogen-containinggas fi 10. along'the path ofihydrocarbon flow through .theafirst Filed 1956 Sm 606386 reactor catalyst. bed in aJparticular manner. According I'Claim. (1; 293-6'5) to my invention. amixturerof hydrogen-containing gas "and hydrocarbona'material ina. molar. ratiolof at :least Thls Invention relates to catalync Rimming of gaso' .2:1, preferably} to'- 4:1, and. at aztemperatureer900 line boiling range hydrqcarbons more Particularly to 975 F. 'is chargedto a catalyst bed in the'first .fixed- {to the catalym reforming of 'stralght'mn or cracked -bed-rea'ctoriof amultiplereactorfixed-bed system. .Ad-

naphtha fractions in the presence of added hydrogen in a ,hydmgenwontainingx gas at temperature *multlplefiagmr y 1-O00 to 1400 F. is separately introduced into' the'first present day icatalytlc treformng' process 18 reactorat a plurality of Points along thepath of hydroiploy'edio upgrade 10W tstraifgiltinin or cracked carbon flow'through th'e'bed. Thepre-heated hydrogeninaphthas' to hlgh octane fractlons'boflmg'm the gimme containing;gas i s intimately admixed at each point ofinnumber of catalyst typqs are used i troduction with the partially reacted hydrocarbonmateq Wlth file most 3 bemg 9 *rial passing-that point. The pre-he'ated hydrogen-con.- ialumma type Convenuonal catalytic rpfonilmg umts ataining gas is introduced at each point along the path' of eiuse' of three more fixed-bid adlabatlc ireiilctofs hydrocarbon' 'flow in stifiicient quantity and is admixed "with .mtgrheatmg between m cider mamtinn with the partially reacted hydrocarbon material passing leacilon temperatures The .basls for @lsgeneta} deslgn "that point in sucha manner that the temperature 'atany ffiges-mi'thefliamre of-the .Ieacuon syste.m Involved m naph' point; along the path of'hy'drocarbon flow within the'cattefmmmg The i catalytlc leactlfmsifire'pre 'alyst' -bd is maintained within about'10 70- u er -the 1dqmmamiylfindotimrmlcand p hydrocarbon material inlet temperature and such that :gemdyihlgh macho rates: cmmsPmdmgly'the sysem the=efiluent gas frorrfi-the first reactor containsa fin'al temmmmI?s3 F Tammy Tand' ltjbectomeil necessary molar ratio of hydrogen contaim'ng gas to hydrocarbon empby mterheatmg thmughwt' thesystem JP mat'erialof'about S' to 15:1; preferably Std 1021. 'The -maintain the desired reaction temperature conditions. total fliuem; fmmYItheQ first reactor is withdrawn at a EEor example, the temperature profile-curves-plottedas ptempeeratumbf85y to, 25o reheatd to tempera f Catalysbbedmveni? reef-90mm 975 and-introduced into a c'at'alyst a typlqalifixed'bed'adlabatlc :three reactor system mdlcite bed in a subsequent fixed-bed reactor. The effluentfrom fi j mfcatalyst bed temperamff of j P 100 a the subsequent reactor is Withdrawn-at -atemperature er i fi react, the "ssow 960 F. "n can-be reheated and further reacted reactonand. 20 .-.-25 F. for the thrrdreactor. Thefirst 4 it be Separamd into a reformam .streampafidja :two reactors havequndesrrable temperature profiiesbe- ;hydmgenicdntining gas streanm VA portion 613 theeghsh pa of the Iapld 1058 m {eacnon temperature The drogen-containing gas stream is'recycled to the'firsti rez-thirdtreactor operates almostisothermallyandcan' there- :acton a p a e I fore be considered satisfactory. This loss in system tem- V My invention ,more fullyxdescribed with reference -.peramre for the first two reactors givesrise to' two nn- Figure 1 pfNaphtha echargel enteringmnefllyis mixed :Kie'sirableconditions: (1)- poor. utllization ofa-catalyst in ,With r i f h ld hydrogen-rich-recycle gas -the. first reactors; and (2) unequal aging of catalyst since Stream. from line 2 mine 3 and preheated to .aging 'proceeds exponentially with temperature i.e., the 1 p in exchafiger mm a Damion ofith x' iflue'nt "higher the Pl f the more rapid! the rate of a n from second reactor 11 passing to exchanger .4 through "The poor utilization of catalyst requires larger catalyst and Y13 eThehot naphthagw ml-xtureis passed inventories and in view of the high costs of platinum type, bymeam: to preheating f m: the p y t becomes a major factvr in deslgn; h temperature-is raised toibetweensoo and 975 "'maqualltles inthe aglng Yates O the catalyst in adlabatlc From furnace 5 the high temperaturezmixtureiis passed I reactors is of considerable importance because of its-efby a 0f li g TialfirsfreagtQrS i m h-ta;p1

on the p Processing y With'the third ;ralitysofEreheatediLZQneS. 'Theeifiuent. from reactor-" 8' reactor in a three reactor adiabaticsystem operating at leaves by ns f line 9 at about 850? to: 92S1F..:a1id -the highest temperature, the catalyst in this reactor will i gd g im h fi fum ifl h i h age most p y and therefore Will fix the length of the :ture; is 'raisedito abont. 900".to 975 F. The efliuentiis .Qp y Since the Processing Cycle Will be c then; passed .Ldown-flowrthrough. second reactor 11, and

P ted before the catalyst in the first two reactors has ta-e second reactor'ei'ffluentiis withdrawn-bymeans of'line aged to an equivalent degreerthis catalyst will not be fully :12: at a temperatureroff-about'880 to960 F. ?The=ereutilize fluent intlinerlzisdivided into two portions passingrre- :In contrast to. such a fixed-bed adiabatic three reactor ,spectively throughrexchangers- 4.and: 15 by means of lines system, a system which operated isothermally'wouldnot 13 and 14 which are recombined in line16 andpassed suffer from these undesirable conditions. Although a through cooler 17 to flash drum 18 maintained at attruly isothermal reactor .systemwould in general not be vmospheric temperature. Flash drum 18.. separates ,athe practical, any approximation to the isothermal'system efiluent intQargas.andliquid-stream.whichlatteris'withwould be desirable. I have now,devised-ainu1ltiple-reac- -drawn .by. means. of line 1-9 and passedto a fractionattirig tor fixed-bed system which approaches isothermal opera- .co1umn (not -shown) -for stabilization. Hydroge'n rich tion, the first'greactor ofwhich replaces-the first two regaspassing -overheadirom-flash drum-18 is divi'ded' into actors otaconventionalthree. or more reactor adiabatic -a net-gas-make-port-ion passing'through line 20 and aremeansfof lines 22 and 23 through gas compressors 24 and 25 prior to being recombined in line 26. The compressed t re'cycle gaspassing through line 26 is divided into a por- -1t-ion -vvhicl1" isg'passed through line 2 and combined with '7 the charge naphtha as desciribed aboveahd a separate portion-which ispassed by line 27 flirough exchanger 7 :15 wherein it"i s preheated to 750 to800 F. prior to *introduction into preheated 28 wherein the temperature is raised to l000 to'l400 Fl The hot gas from furnace v f f28' is passed'by'meansof lines 29, 30, 31, 32 and 33 to "ifirst reactor 8. Each of lines 29, V30, 31, 32 'and'33 is provided with'a 'separatetemperature control system so, that the amount'of heating can be directly regulated; t The general features of the design of first-reactor '8 areshown iu Figures 2-and 3'; This. reactor is divided p into l6'catalytic zones, a, b, c, d, e and 1'', each zone being separatedby means of gas distribution sections 35, 36, V 37, ,38' and 39. These gas distribution sections can be composed of some" catalytic'ally, inert. material such' as 'alumina spheres, granular Alundum, 'or the like." Situ- .ated;?in the respective sections 35-39 are distribution -heaciers40, 41, 42, 43 and provided with ,a'large'numberfof 'perfOrationsfl45' shown in detail in Figure 3.1

7' a The hot naphtha-gas mixture-fromlin'e 7 enters reaca mtor 8, is passeddown-flow through :the upper-catalyst re- -;tentiqn;zone:into the top catalyst zone a. The partially -f-; r,eacted fnaphthagandjrecyclefgas leave thisifirst catalyst me ma flow zinto first, reheat distribution section 35 a yvherein; the naphthagas stream is v intimately mixed with the 'high'temperature recyclejgas. entering through'line 291- and header- 40;; -The reheated gas mixture is then ereheating-isr repeated in each 'offsections;3610'39.

W q'lhe amountjofi reheating gas introduced into eachof sectionsSS-"IBv is'regulatedby'meansof'a control valve 7 located zi'n the corresponding incominggas line 29-33 g 'whichvalve .isg'responsive to a temperature controller 7 maintaining the exit temperature from the catalyst zone l involved. *Thus the degree of reheat for each zone can i -bea 'ya iu d-- 11 r i ln-Figurel; five reheat sections have been, shown and a the amount of catalyst requiredlfor each zone is'shown substantially in a proportion Jwhich-gives satisfactory operation; 'The increase the amountof catalyst re-- ,quired after eachreheat section is the result ofithe lowered-reaction; rate as the reaction proceeds toward Lequ'ilibrium; The number and arrangem'ent of the cat-' alyst zones and reheat sections, however, can be varied l to suit any, particular naphtha type'or temperature pro- 7 Q ,file required Also -the reactor shdwnin Figure 2 is p I subject to manyotherdesign arrangements'such as Vania c i; tions in the design features of the jcatalyst beds. VFo'r ;example, aecross-flow or;annular type reactor can'be V iiused. Variations in the manner by which the reheat mix- Zingis achievedcan alsobe made;

Figures land izgasre'ontainin g 'about 70'volumes percent hydrogenfat,

$9 205 Fnfrom furnace 15 is passed by means of line 7- to V cylindrical reactor '8 atan inlef'pressure of 500 psig asseg-into catalyst-zone'b fof'thecfurther reaction and V 7' Additionalrecycle gas from furnace 28 at l140 F. is

passed by means of'lines 29, 30, 31, 32 and 33, into distribution headers 40, 41, 42,43 and respectively in f the following amounts: 2740 mols/hrg 36 50 mols/hri; 4560 mols/hr.; 5500 m'ols/ hrg and 6400 mols/hr. The vapor exit temperatures from catalyst zones a through f pounds of catalyst.

' with an overall space velocity, of 2.0 WHSV, the catalyst";

' perature, it is the third reactor'th-at fixes the naphtha .The efiiuentfrom reactor Sata temperature of 868 F. is passed by means of line 9 to interheating furnace 10 wherein the temperature is raised to 920 F. It is introduced into reactor 11 at an inlet pressure of 500 p.s.i.g. Reactor 11 is cylindrical and contains 18,500

Eifiuent from reactor 11 is withdrawnat90l F. a 7 V Reformate is separated from the efifluent in 82% yield of C and has an octane number of 95 RON i 1 v i For a conventional adiabatic system involving} equal sized reactors operating with equal inlet temperatures and inventory in each reactor is about 18,500 pounds. Sinc e reactor, and since catalyst'agingis' a function oftentthe highest average bed temperatur'eobtains in thethird 1 n processing cycle time (before regeneration or replace-V V ment'land under suchiconditions the catalyst in the first two reactor" beds has; not been effectively utilized Although the -operation of second reactor 11' in the above illustration is essentially the's-anie" as thei operation of the third reactor in a conventional adiabaticsystem as V can'be seen from the above illustrativeoperafiom -the', 7 "process of my invention affords significant'savings in V a catalyst" inventory provides more effective tutiliza tion of a larger proportion of the catalyst inventory;

For example, inlthe illustrative operation catalyst vessels containing 37,000"pounds-of catalyst are 'replaced: by one vessel containing 125825 pounds of 'catalystg'jl Certain reactor conditions in the pro'cess'of my in-w venti-on can, of course, be varied. For example, there a'ctorpressureican rah-gejfrom about 'toiabout 11000 v I V 'l'I'he eatalystinventory in reactors, whichcomprisesa 7 conventional pelleted platinum on alumina catalyst, is divided'as'followm a r pLs.i.g. and the weight hourly. space velocity can range 1 from about 0.25 to 10.0. Also the'proportion of hydro- -genjin .thenrecycle gasjcan'range from about 35% to about 95% byvolume. T

-:A determination of thepro'portions of hot hydrogen in containing recycle'gas separately introduced into the first ,or semi-isothermal, reactor and thelproportionsfof cat-a: 'lyst in each of the catalyst zones in the semi-isothermal reactor can be c'onvenientlymade by referencelto Figures 4and Sand tothe calculations appearing below which they are derived.

One molenaphthene one r nole aromatics expression:

P,,=partial pressure 'of naphthenes insystem P,,=partial pressure of aromatics in system ii (B =partial pressure of: hydrogen in' system,

7 a V V n w m q c Thisrelationship can also be represented M me general 'equation tori the equilibrium constant of the above;

r y 7- w 7 H First a heat balance is developed relating thefheat quantities involved for the recycle gas 'andenaphthafeed streams and the primary system reaction; t at ofpdehy drogenation of naphthenes; toaromatics; a From such V .a heat balance study, it is possibleto relate reaction tem -V V peratures to the, degree ci -reaction p show'n byethefireli ,lationship:

Equation 2 can also be rewritten in terms of the percentages of naphthenes-aromatics and total system:pressure:

K'=[(percent A)/ (percent N)] x (partial. (3)

pressure of hydrogen) This equation indicates that the ratio of'aromatics to naphthenes is a function of'the partial pressureof the hydrogen present. It is, therefore,.possible todevelop a series of relationships between reactiontemperature and the aromatic-naphthene ratio and various partial pressures of hydrogen. These relationshipsare plotted in Figure 4 as the change in naphthenecontent (naph thenes infeed minus naphthenes in product) versus'rewhen temperature. For convenience, the hydrogen partial pressure parameters have been termed recycle gas ratios, i.e., 2/1-3/1, etc.

The conversion of naphthenes to aromatics can be shown by a heat balance to be represented by the relationship AT=CAN 4 in which AN=change in naphthene content (feed minus product) AT=change in temperature C=constant where the. constant C is a function of the quantity of recycle gas (e.g., partial presssure of hydrogen). With such a relationship a series of .operating lines for various recycle ratios can be plotted as shown in Figure 4. These operating lines have an origin based upon a heat balance. Therefore, if a fixed quantity or" heat is added to a fixed quantity of recycle gas, this is equivalent to increasing the temperature of the inlet feed system (naphtha plus recycle gas). For illustration, it is assumed that normal operation would be represented by a gas recycle ratio of 7.0 mols of gas per mol of naphtha and that the semi-isothermal system would be represented by operation with multiple reheat steps. Thus the initial operation would be at a 920 F. inlet temperature with a recycle ratio of 2/ 1 mols per mol, and each subsequent step would be with an additional unit of recycle gas as shown. By firing the temperature of the recycle gas, each subsequent operating line can be located. It will be noted that each of these operating lines terminates on a different equilibrium line.

This graphical representation of the reaction system thus allows for a ready evaluation of the stages of reheat. For the case shown in Figure 4, the reaction is initiated at a temperature of 920 F. and proceeds along the 2/ 1 recycle ratio operating line to a temperature of 894 F. at which point the first reheat is used to raise the system reaciton temperature to 920 F. The reaction then proceeds along the 3/ 1 recycle ratio operating line to a temperature of 891 F. at which time the second reheat is started. These steps are tabulated below:

The graphical representation of conventional adiabatic operation is shown by a dotted line on Figure 4 with the initial operation at a temperature of 920 F. and a recycle ratio of 7/ 1. This is allowed to proceed to a temperature of 803 F. (A) which is the outlet of the No. l reactor; at this point the reaction products are reheated to 920 F.

(B) with the interheater .and the reaction allowedto' pm ce'edlto a temperature of 870 F. (C), the outlet of the No. 2 reactor.

It is evident that the system chosen representsthere placement of the first two conventional adiabatic reactors by oneequivalent semi-isothermal reactor. It is, however, apparentthat there are a limitless number of 'com-. binations of reheat spaces, temperature combin'ations'and recycle ratios that could be chosen.

Having once fixed'the reaction system, it is necessary to determine the amount of'catalyst needed inieach of the reheat stages above. This is done by meansofFigure 5 which is a plot of the reciprocal space velocity'with an expression relating inlet and equilibrium temperatures with the temperature at any point in the system.

The amount of catalyst calculated .for each of the .severalreheat stages is based upon the amountof catalyst required for a comparable adiabatic system. However, since each of the reheat stages is operating with different conditons of equilibrium temperature and exit tempera.- ture for the reheat stage under consideration (when compared with the comparable adiabatic system), it is'necessary to make a suitable adjustment in the catalyst requirement to reflect these new temperature conditions. This adjustment is made by means of the following relationship:

K g- /RT 5 This can be simplified (1 ln K E 2 d T A /RT Assuming AE constant an AB r K R T T If IC =1 at a fixed temperature (exampl e=920 F.)

AB 1 1 111 KT2R or -ar 1 1 R KTF T. a/ (9) in which Substituting the inlet and equilibrium temperatures for the various reheat stages, values of the rate constants are calculated for each reheat stage. These values are then used to adjust the space velocity requirements to the new conditions.

'The development and application of this procedure in- Y volves several assumptions which should be considered.

(1) For simplification, the reaction system used for the first two reactors considers only the predominating naph- Lhene dehydrogenation reactions. There are other reactions involved to a lesser extent, however, but these have been neglected. Possible errors from this source are not of significance, although for exact cases consideration should be given to such reactions.

(2) The heat balances usedhavebeen based upon the assumption of constant reactor product distribution atall points in the system. Actually," however, the product dis-.

, tribution is subject to continual change as the reaction ata plurality of points along the path of hydrocarbon proceeds through the catalyst bed. This assumption intro duces" an error of only small magnitude and can, there fore, normally beneglected. t

(3) Thefdetermination of theamount of'catalyst re quired in each ofthe, reheat stages (see Figure 4) is based upon the assumption of perfect mixing in each of the reheat zonesu 'Since perfect mixing and subsequent vapor distributionis'diflicult to insure, it is probable that a small amount of additonal catalyst would be required in each flow through the bed, intimately admixing the pre-heated hydrogen-containing gas at each point of introduction with the partially reacted hydrocarbon material passing that point, the pre-heated hydrogen-containing .gas being introduced zit-each point along the pathof hydrocarbon llow in. progressively increasing quantities and being adrnixed in such a manner that the temperature at any point alongthe path of hydrocarbon flow withinthe, catalyst 1 bed is maintained within about lO-70 F. of the hydroe carbon material inlet temperature and such that the emu-,

ent gas from the first reactor contains a final molar ratio of hydrogen-containing gas to hydrocarbon material of about 5l5 :1, withdrawing the total eflluent from the first reactor at a temperature of 850 to 925 F., reheating the effluent from the first reactor to a temperature of system which comprises charging a mixture of hydrogen- I I containing gas and hydrocarbon material in a molar ratio of at least 2:1 and at a temperature of 900 to 975 F.

'to a firstfixed-b ed' reactor having a plurality of catalyst zones containing platinum;on-alumina type reforming catalyst, said zones in the course of the hydrocarbon flow containing a progressively larger amount of catalyst separately introducing additional hydrogen-containing gas at a temperature of 1000" to 1400 F. into the first reactor 900 to" 975 F., introducing the first reactor etfiuent to a catalyst bed ina subsequent fixed-bed reactor, and with drawing the efiiluent from the subsequent reactor at a temperature of 880 to 960 F.

References Cited in the file of this patent UNITED STATES PATENTS 2,330,069 Marshall Sept. 21,1943 2,349,045 Layng et al. Mayj16, 1944 2,418,534 Watson Apr; 8, 1947 2,439,934 Johnson et al. Apr. 20,1948 2,759,876 Teter et al. Aug; 21, 1956' 'l- Iaensel et a1 Feb. 12; 1957 

